Process for producing butyl acrylate

ABSTRACT

An improved process for producing n-butyl acrylate in high yield and high purity substantially free of acrylic acid, incorporates one or more of the following new process components in an acid-catalyzed esterification process for producing n-butyl acrylate: 
     1. A hydrolytic recovery component, wherein heavy end adducts produced during the acid-catalyzed esterification are hydrolyzed, recovered, and recycled as valuable reactants from a hydrolytic recovery unit (HRU); 
     2. A cracking reactor component, preferably used with the HRU, wherein additional valuable reactants are recovered and recycled after treatment in the cracking reactor; and 
     3. A new distillative component, wherein a crude n-butyl acrylate stream is efficiently distilled in an aqueous mode through an acrylic acid separation column, thereby providing n-butyl acrylate substantially free of acrylic acid and in high yield. 
     The first two components also are applicable to acid-catalyzed processes producing C 1  to C 4  alkyl acrylates. A continuous process producing n-butyl acrylate incorporating all new process components also is disclosed.

This application is a continuation of U.S. patent application Ser. No.09/049,483 filed Mar. 27, 1998, now U.S. Pat. No. 5,990,343 which is acontinuation of U.S. patent application Ser. No. 08/797,380 filed Feb.7, 1997, now U.S. Pat. No. 5,877,345.

The present invention relates to an improved process for producing butylacrylate. More specifically, the invention relates to a new method ofdistilling, and of recovering and recycling normal butanol (“BuOH”),acrylic acid (“AA”), and normal butyl acrylate (“BA”) from one or moreprocess streams in an acid-catalyzed esterification process for BA. Theinvention encompasses two new process components, one related to thehydrolytic recovery of valuable reactants from their higher boilingadducts, and a second component related to improved distillation of acrude product yielding BA substantially free of AA. The hydrolyticrecovery component of the invention also is useful in processes forproducing selected acrylic esters, in addition to BA. Most specifically,the invention relates to a highly efficient, continuous process forproducing BA in high purity and high yield.

Direct esterification of AA with an alcohol is an equilibrium process.The equilibrium constant determines the net rate and extent ofconversion of AA and alcohol; for continued high rates of conversion themixture must not approach equilibrium. Conventionally, an excess ofalcohol over AA is employed and water of esterification is removeddistillatively as its azeotrope with alcohol and ester to maintain ahigh rate of conversion of AA. The azeotrope is removed via adistillation column mounted directly on an esterification reactor. Inthe case of methyl or ethyl esters, the water of esterification, excessalcohol, and product ester are removed from the head of the distillationcolumn and are substantially free of AA. Water extraction removes thealcohol which is concentrated distillatively for recycle to the reactor.The washed ester is azeotropically dehydrated and finally distilled toprovide the pure ester product. In butyl acrylate production, however,the separation of acrylic acid from water of reaction, excess alcohol,and product ester is more difficult, and the distillate from theesterification reactor in a continuous process typically contains 1-3%AA. This AA typically is extracted into aqueous caustic. Although it ispossible to recover some of this AA from the resulting aqueous saltsolution by acidification with a strong acid followed by extraction intoan organic solvent, e.g. butyl acrylate or butyl acrylate/butanolmixture, significant loss to a large aqueous waste stream isunavoidable. The butyl acrylate and excess butanol are nextazeotropically dehydrated wherein excess butanol is separated from theproduct ester as a butanol/butyl acrylate azeotrope for recycle to theesterification reaction. A final distillation provides pure butylacrylate. In all cases a small bleed stream is removed from theesterification reactor and a small bottoms stream is taken from thefinal product distillation to remove high boiling byproducts andinhibitor residues from the process. These streams are stripped torecover free AA, alcohol, and alkyl acrylate values, but little or noneof the values present within the high boiling byproducts are recovered.Thus, the conventional processes for producing C₁-C₄ esters suffer fromyield losses to high boiling byproducts, and the C₄ process furthersuffers from direct losses of AA because of the difficulty in separatingAA from butanol, water, and ester.

In the art of recovering and recycling reactants from their higherboiling adducts formed during processing (so called “heavy ends;” in BAproduction these include, for example, butyl β-butoxypropionate andesters of sulfuric acid), there has been only limited success. Forexample, in ethyl acrylate (“EA”) production from ethylene and AA, U.S.Pat. No. 4,968,834 ('834) describes a process for recovering EA from a“spent black acid” stream containing sulfuric acid residues and otheradducts bled from the bottom of a distillation column. The '834 processuses an alcoholic solvent to facilitate an overhead distillativerecovery of ethyl acrylate, and treats the black acid residues with anaqueous alkanol mixture. No materials are directly returned to theEA-producing reactor nor to the distillation column which generates theblack acid stream. The '834 process thus provides partial recovery ofethanol, EA and AA, but only by an aqueous treatment which is isolatedfrom the reactor of the ethylene-AA process. Other processes employdistillation units (often designated “bleed strippers”) to partiallyrecover free AA, BA, and BuOH from reaction bleeds, but to the extentthat heavy ends are recovered in that operation, they remain chemicallyin the higher boiling (heavy end) form and are not transformed to thedesired valuable AA, BA, and BuOH forms.

Distillation is commonly used in BA production. For example, U.S. Pat.No. 4,012,439 ('439) describes a continuous process for BA in which areactor esterification mixture is distilled through an AA separationcolumn to give an overhead mixture of BA, butanol, and water, and, fromthe column bottom, a concentrated AA stream which is returned to thereactor. While separating the overhead mixture from AA, the '439 processrecycles a very high proportion (>97%) of aqueous phase distillate tothe head of the AA-separating column. This high proportion of aqueousrecycle (i.e. having an aqueous reflux ratio of about 32:1)disadvantageously requires a large column and a large expenditure ofenergy in returning large volumes of water to the process.

Thus, in the acid-catalyzed production of acrylic acid alkyl esters(“alkyl acrylates”), particularly of BA, there remain significant energyuse and reactant recovery problems. There are needs for a process whichwould recover reactants from their higher boiling, heavy end, adductsformed during the production of acrylic esters, e.g. BA, which wouldrecycle recovered reactants and the ester to the esterification reactoror elsewhere in the process for reuse. Further needs include methodsmaking more efficient use of the water of reaction, both in facilitatingdistillative separation of acrylic ester from AA and in more efficientlyrecovering and recycling unreacted AA, particularly if these steps wereaccomplished with reduced energy use. Meeting one or more of these needswould provide increases in process and/or material use efficiencies.Additionally, if such improved processes led to reduced dibutyl ether(DBE) byproduct in comparison to known processes, even greater processefficiency would result.

We have discovered a high yield process for producing alkyl acrylates,using BA as a preferred example, which achieves these desirable ends.Our new process provides for the recovery of “values,” that is,reactants and alkyl acrylate product, from the heavy ends produced inthe process. Our new process includes the use of at least one of thefollowing process components: 1. recovering values from a hydrolysisreactor unit (“HRU”) fed with a source of heavy ends, as from anesterification reactor; 2. recovering additional values from a crackingreactor preferably used in conjunction with the hydrolysis reactor; and3. specific to a continuous BA process, distilling by use of an acrylicacid separation column in an efficient new way and providing recovery ofBA which is substantially free of AA. Our new process advantageouslyprovides very low levels of DBE in product BA because the esterificationreactor is operated under mild temperature and pressure conditions, andat relatively low acid catalyst levels.

Thus, in the broadest use of the hydrolytic recovery component of theinvention, there is provided a method of recovering AA, a C₁-C₄ alkylacrylate, and a C₁-C₄ alkanol from heavy ends produced during productionof the C₁-C₄ alkyl acrylate, comprising the steps of:

a) feeding a total aqueous and heavy end feed stream comprising theheavy ends, water, residual acid catalyst, and optionally a strong acidselected from a mineral acid or sulfonic acid, to a hydrolysis reactormaintained at 90° to 140° C., 50 to 1000 mm Hg pressure, and a residencetime of 0.5 to 20 hours based on the total aqueous and organic feedstream;

b) distilling an overhead stream containing AA, the C₁-C₄ alkylacrylate, the C₁-C₄ alkanol, and water from the hydrolysis reactor whilemaintaining a hydrolysis reactor liquid concentration of from 5 to 40weight % water and at least 1 weight % acid, the acid comprising theresidual acid catalyst and the optional strong acid;

c) condensing the overhead stream;

d) separating from the condensed overhead stream an organic phasecomprising the C₁-C₄ alkyl acrylate, the C₁-C₄ alkanol, and AA, and anaqueous phase comprising primarily water, and AA and the C₁-C₄ alkanol;

e) removing the separated organic phase;

f) recycling the separated aqueous phase to the hydrolysis reactor; and

g) withdrawing from the hydrolysis reactor from 20 to 70 weight %, basedon the total aqueous and heavy end feed stream, of a hydrolysis reactorbleed stream.

Specific to BA production, there is provided a method of recovering AA,n-butyl acrylate (BA), and n-butanol (BuOH) from heavy ends producedduring acid-catalyzed esterification of AA with BuOH, comprising thesteps of:

a) feeding a total aqueous and heavy end feed stream comprising AA, BA,BuOH, water, heavy ends, residual acid catalyst, and optionally a strongacid selected from a mineral acid or sulfonic acid, to a hydrolysisreactor maintained at 90° to 140° C., 50 to 1000 mm Hg pressure, and aresidence time of 0.5 to 20.0 hours based on the total aqueous and heavyend feed stream;

b) distilling an overhead stream containing AA, BA, BuOH, and water fromthe hydrolysis reactor while maintaining a hydrolysis reactor liquidconcentration of from 5 to 40 weight % water and at least 1 weight %acid, the acid comprising the residual acid catalyst and the optionalstrong acid;

c) condensing the overhead stream;

d) separating from the condensed overhead stream an organic phasecomprising the BA, the BuOH, and AA, and an aqueous phase comprisingprimarily water, and AA, and BuOH;

e) removing the separated organic phase;

f) recycling the separated aqueous phase to the hydrolysis reactor; and

g) withdrawing from the hydrolysis reactor from 20 to 70 weight %, basedon the total aqueous and heavy end feed stream, of a hydrolysis reactorbleed stream.

Another embodiment of the invention provides a method of continuouslyrecovering AA, n-butyl acrylate (BA), and n-butanol (BuOH) from heavyends produced during acid-catalyzed esterification of AA with BuOH,comprising the steps of:

a) withdrawing continuously a reactor bleed stream from anesterification reactor containing an esterification reaction mixturecomprising AA, BA, BuOH, water, heavy ends, and residual acid catalyst,while concurrently distilling AA, BA, BuOH, and water from theesterification reaction mixture;

b) feeding a total aqueous and organic feed stream comprising thereactor bleed stream, water, optionally a strong acid selected from amineral acid or sulfonic acid, and optionally additional heavy ends, toa hydrolysis reactor maintained at 90° to 140° C., 50 to 1000 mm Hgpressure, and a residence time of 0.5 to 20 hours based on the totalaqueous and organic feed stream;

c) distilling an overhead stream containing AA, BA, BuOH, and water fromthe hydrolysis reactor while maintaining a hydrolysis reactor liquidconcentration of from 5 to 40 weight % water and at least 1 weight %acid, the acid comprising the residual acid catalyst and the optionalstrong acid;

d) condensing the overhead stream;

e) separating from the condensed overhead stream an organic phasecomprising BA, BuOH, and AA, and an aqueous phase comprising primarilywater, and AA, and BuOH;

f) removing the separated organic phase;

g) recycling the separated aqueous phase to the hydrolysis reactor; and

h) withdrawing from the hydrolysis reactor from 20 to 70 weight %, basedon the total aqueous and organic feed stream, of a hydrolysis reactorbleed stream.

Additional recovery of valuable reactants is achieved from heavy ends byusing a cracking reactor in tandem with the hydrolytic recovery methodsdescribed above. That process is carried out with any of theabove-described hydrolytic recovery methods by further including thesteps of:

a) feeding up to 100% of the hydrolysis reactor bleed stream to acracking reactor maintained at 90° to 140° C., a pressure of from 20 to200 mm Hg, and a residence time of 0.5 to 20 hours based on the fedreactor bleed stream;

b) distilling from the cracking reactor a cracking reactor overheadstream comprising AA, the C₁-C₄ alkyl acrylate, the C₁-C₄ alkanol, andwater while maintaining a cracking reactor liquid concentration of atleast 7.5 weight % acid;

c) condensing the cracking reactor overhead stream; and

d) recovering from the cracking reactor overhead stream AA, C₁-C₄ alkylacrylate, C₁-C₄ alkanol, and water.

Preferably, the alkyl acrylate is BA. More preferably, the crackingreactor just described is used in tandem with the hydrolytic reactor ina continuous acid-catalyzed process for producing BA.

In a second component of the invention, this relating to continuousproduction of BA, there is provided a method of continuously recoveringn-butyl acrylate (BA) substantially free of AA from an esterificationreaction mixture, comprising the steps of:

a) feeding continuously to an esterification reactor AA and BuOH in amolar ratio of from 1 to 1.1 to 1 to 1.7, and an acid catalyst;

b) reacting the AA and BuOH to yield BA in a conversion of at least 60%on AA, and yielding the esterification reaction mixture comprising AA,BA, BuOH, water, heavy ends, and acid catalyst;

c) distilling from the esterification reactor a vaporized mixturecomprising AA, BA, BuOH, and water;

d) condensing the vaporized mixture to provide a first condensatecomprising an organic phase and an aqueous phase;

e) returning from 0 to 30 percent of the organic phase to an entrainmentseparator surmounting the esterification reactor; and

f) feeding from 70 to 100 percent of the organic phase and from 50 to100 percent of the aqueous phase to an acrylic acid separation column;

g) distilling from the acrylic acid separation column, at a pressure offrom 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratioof 8.5:1 to 17:1, an overhead mixture comprising an azeotropic mixtureof butanol, butyl acrylate and water;

h) removing from the distillation column an acrylic acid-rich bottomstream;

i) recycling the acrylic acid-rich bottom stream from the acrylic acidseparation column to the esterification reactor;

j) condensing the overhead mixture to provide a second condensate;

k) separating the second condensate into a butyl acrylate-rich organicphase and a separated aqueous phase; and

l) removing the butyl acrylate-rich organic phase substantially free ofAA.

The recovering of BA substantially free of AA also may be carried out byfeeding the vaporized reactor mixture directly to the AA separationcolumn by bypassing the d), e), and f) steps immediately above. When thevaporized mixture is fed directly to the column, the aqueous refluxratio is tightened to 13:1 to 17:1; all other steps are identical,except there is, of course, no “first condensate.”

BRIEF DESCRIPTION OF THE DRAWINGS

In a brief description of the drawings, the process embodying bothcomponents of the invention is shown schematically in FIG. 1; FIG. 2 isa graph of the amount of residual AA in organic distillate obtained bydistilling through the acid separation column versus the aqueous refluxflow rate, as obtained under conditions described below. FIG. 1 showsequipment and flow lines, including esterification reactor 1, bleed line3 to a hydrolysis reactor unit (HRU) 5 ; and associated streams andlines, particularly line 8 returning organic phase to reactor 1 and line7 returning aqueous phase to the HRU. Cracking reactor 10 also hasassociated lines, e.g. for draining and distilling, and providing forreturning the condensed overhead stream from separator 31 to reactor 1via line 12. In a brief description of the drawing relating to thedistillation component, FIG. 1 includes line 2 feeding an esterificationreactor vaporized mixture (in this embodiment) to a condenser 62 and thecondensate to a phase separator 14 and associated lines from the phaseseparator to the acrylic acid separation column 15 via one or more line43, 53, and for optionally feeding some of the aqueous phase to the HRU5 via line 42 and optionally a portion of the organic phase to anentrainment separator surmounting reactor 1 via line 41, when separator14 is used. Line 54 provides for optional feed of BuOH to the AAseparation column. Lines from the acrylic acid separation column include17, returning the AA-rich bottom stream to reactor 1, and line 16conveying the distilled overhead mixture through condenser 63 to phaseseparator 18 and its associated lines, line 21 returning a controlledportion of the aqueous phase 20 to the top of acrylic acid separationcolumn 15, line 22 moving forward a controlled portion of the separatedaqueous phase and line 23 carrying forward all of the BA-rich organicphase, 19. Line 52 provides for BA/BuOH return to reactor 1 duringsubsequent conventional processing and final BA product isolation. FIGS.1 and 2 are described in greater detail below.

DETAILED DESCRIPTION OF THE INVENTION

The first component of the invention, the hydrolytic recovery componentwhich recovers values from heavy ends, takes advantage of the knownability of a strong acid, for example, a mineral acid such as sulfuricacid, to catalyze the individual reactions employed: directesterification, ester hydrolysis, dehydration, and retro-Michaelreactions. Thus, catalytic processes in which esters and heavy endhydrolyses occur in a hydrolysis reactor and, in an extended embodiment,dehydration and retro-Michael additions carried out in a crackingreactor, are new efficient methods for recovering, for example, BA,BuOH, and AA values from heavy end components formed during priorreaction in, for this example, a BA esterification reactor. The heavyends are exemplified in detail for BA production using sulfuric acidcatalyst in the reactor and in the HRU; from these examples one skilledin the art would recognize analogous “heavy end” counterparts fromproducing any of the C₁-C₄ alkyl acrylates. The C₁-C₄ alkyl groups maybe methyl, ethyl, propyl and iso-propyl, and the butyl isomers,preferably n-butyl. Heavy ends are adducts higher in boiling point thanthe reactants and, as exemplified here, the butyl acrylate product; theyinclude, for example, acryloxy-propionic acid (“AOPA”) and its butylester derivative, beta-hydroxy propionic acid and its butyl esterderivative, beta-butoxy propionic acid and its butyl ester derivative,and other non-polymeric adducts of the reactants. In addition, maleicacid and benzoic acid impurities in the acrylic acid and sulfuric acidcatalyst are present as maleic acid monobutyl ester, butyl benzoate, andmono-butyl sulfate. Furthermore, the simultaneous removal of BA, BuOH,and AA by way of the distillate streams of both hydrolysis and crackingreactors in a continuous process allows the recovery reactions toproceed beyond equilibrium constraints present in a batch process andthus improve process yields. Another advantage of the hydrolyticrecovery component of the invention is that one or more additional heavyend streams may be worked into the hydrolytic recovery process stream,thus providing recovery of additional values.

The following are examples of heavy end materials present in the totalaqueous and organic (i.e. heavy ends alone, or mixture of heavy ends,reactants, and product) feed stream which are hydrolyzed in thehydrolysis reactor to afford valuable recoveries of AA and the describedalkyl acrylates and alkanols. Alkyl esters of the β-alkyloxy propionatesare a common heavy end material. In the beta position of the alkylesters also may be the hydroxy group instead of an alkyloxy group.β-Acryloxy acid derivatives of the C₁-C₄ alkyl esters also may bepresent in the heavy ends; for example, butyl (β-acryloxy) propionate iscommonly present in the heavy end materials along with its correspondingacid, in BA production. Also present are the C₁-C₄ esters of sulfuricacid catalyst which esters are hydrolyzed to sulfuric acid and thecorresponding C₁-C₄ alkanol.

The reactions which take place in the HRU can be generalized byequations 1 and 2, following:

Here, R¹ is a C₁-C₄ alkyl group as defined above; R² is a C₁-C₄ alkylgroup or H. Additionally, saturated and unsaturated esters, such as theC₁-C₄ alkyl ester of benzoic acid and the C₁-C₄ alkyl ester of maleicacid as well as C₁-C₄ alkyl sulfate, can be similarly hydrolyzed torelease an equivalent of C₁-C₄ alkanol. Furthermore, the simultaneousremoval of BA, BuOH, and AA via the HRU's distillate stream allows therecovery reactions to proceed beyond equilibrium constraints and improveoverall process yields. However, the parent carboxylic acid of severalheavy materials cannot be recovered in the HRU and, therefore, anadditional recovery scheme is necessary for these materials and iscarried out in a cracking reactor.

The reactions which take place in the cracking reactor can begeneralized by equations 3 and 4, following, where R² is as definedabove:

The conversion of C₁-C₄ alkyl ester of beta-hydroxy propionic acid,beta-alkoxy propionic acid, and beta-acryloxy propionic acid to theparent acid in the HRU (via ester hydrolysis) is quite beneficial sinceit is well known that these materials undergo dehydration andretro-Michael addition in acid form. Therefore, under the relatively dryconditions of the cracking reactor, compounds such as beta-n-butoxypropionic acid and beta-hydroxy propionic acid can undergo dehydrationresulting in recovery of acrylic acid and butanol. The well known dimerof acrylic acid (AOPA) undergoes cracking to yield 2 moles of acrylicacid. Here again, the continuous removal of products allows thereactions to proceed beyond equilibrium constraints and improve theoverall process yield.

Referring to FIG. 1 in the working of the preferred continuoushydrolytic recovery component of the invention in producing BA:esterification reactor bleed stream 3 is fed from esterification reactor1 to hydrolysis reactor unit (“HRU”) 5. The hydrolytic recovery methodof any embodiment of the invention may be carried out in a multiplatereactive distillation column or other staged reactor, and preferably iscarried out under continuously mixed conditions, as in a continuous flowstirred tank reactor (“CSTR”). By “bleed stream” is meant any processstream which is controllably withdrawn from one vessel to another, suchas from a reactor to another reactor or distillation column. Here, theesterification reactor bleed stream 3 contains acid catalyst, water, AA,BA, BuOH, and heavy ends; polymerization inhibitors also may be present.Additional feeds via 4 may include water, and may also include mineralacid, for example, sulfuric acid, or a sulfonic acid such as methane-,benzene-, or toluene-sulfonic acid. The mineral or sulfonic acid isadded as required to meet the specified minimal usage level in the HRU.One or more additional streams containing heavy ends from sources otherthan the esterification reactor also may be added. These feeds may beadded by one or more feed lines represented by 4. The additional heavyends may comprise up to 80 wt. % of the total aqueous and organic feedstream. Sulfuric acid is most preferred for use as both reactor acidcatalyst and mineral acid in all embodiments of the invention. The HRUmixture of the described feed streams is maintained in a boiling stateunder the conditions defined. The residence time of from 0.5 to 20 hoursis based on the total aqueous and organic feed stream (“total” meaningthe sum of the aqueous and heavy ends and/or reactor bleed streams) fedto the HRU. Preferred residence time is from 0.5 to 5 hrs, and morepreferred is 0.5 to 3 hrs. An overhead stream is distilled from the HRUmixture in 6 and condensed, 60, into phase separator 30. The condensedoverhead stream separates into an organic phase rich in BA, BuOH and AA,and into an aqueous phase containing primarily (i.e. >50%) water andsome BuOH and AA. The separated aqueous phase is returned to thehydrolysis reactor 5 via 7 and the separated organic phase, as stream 8,is returned in this embodiment to the esterification reactor 1, thusrecovering valuable BA, BuOH, and AA for subsequent reaction and productrecovery. The separated organic phase also may be fed to separator 14for recovery by way of line 43 and distillation through 15. Anundistilled residue, from 20 to 70 wt. % of the total aqueous andorganic feed stream, is bled as the hydrolysis reactor bleed stream 9from the hydrolysis reactor for further handling (e.g. as a waste streamby line 51 or, preferably, as feed to a cracking reactor by line 9.)

A preferred embodiment provides additional recovery of AA, BA and BuOH.As shown in FIG. 1, the hydrolysis reactor bleed stream 9 is fed to acracking reactor 10 and treated as now described. The cracking reactormay be of construction similar to that of the HRU and is, preferably, aCSTR. The cracking reactor liquid is maintained at least at 7.5 wt. %mineral acid, preferably sulfuric, and also contains a mixture ofacrylic acid, BuOH, BA, some heavy ends and residual polymerizationinhibitors. Additional mineral or sulfonic acid may be added to thecracking reactor liquid (feed line not shown). The cracking reactormixture is maintained in a boiling state under the previously describedcracking conditions while an overhead stream is distilled from thecracking reactor via line 11 and condensed via 61 to separator 31. Thecondensate contains an organic distillate stream containing AA, BA andBuOH, and also some water; all of the condensed overhead stream isreturned as stream 12 to the esterification reactor, thus providingadditional recovery of valuable AA, BA, and BuOH. The cracking reactorresidue stream 13 is drained off for further handling, generally aswaste. Preferred and more preferred cracking reactor residence times arethe same as described for the HRU, namely, 0.5-5 hrs and 0.5-3 hrs,respectively.

The HRU may be a multi-plate reactive distillation column so long assufficient number of plates are incorporated to provide specifiedresidence time. When a reactive distillation column is employed as anHRU, a separate cracking reactor unit may not be needed to achieveacceptable values recoveries. Under most production conditions it ispreferred to use the cracking reactor in tandem with a hydrolyticreactive distillation column, similar to its use when the HRU is a CSTR.One disadvantage of a reactive distillation column over a CSTR is thatoccasional build-up of solids on the column trays may requireundesirable down time for column cleaning.

Addition of one or more additional feed streams to the esterificationreactor bleed stream or directly to the hydrolysis reactor permitsadditional recovery of AA, and for example, BA, and BuOH, through theprocesses occurring in the hydrolysis reactor and, when used, thecracking reactor. The liquid in the hydrolysis reactor has at least 5wt. % water for efficient operation; preferably the HRU liquid containsfrom 9 to 18 wt. %, more preferably from 10 to 16 wt. % water, in orderto achieve efficient hydrolysis rates under nominal thermal and pressureconditions and practical equipment size. Water content is maintained bya combination of returning the entire condensed and separated aqueousstream in line 7 to the hydrolysis reactor and by adding additionalwater from other sources, e.g. by lines 4 and 42, to compensate forwater losses to organic distillate and the HRU bleed stream. Wateraddition from the distilled aqueous phase from the esterificationreactor, by line 42, is a preferred source of water in the continuous BAprocess. In order to maintain efficient dehydration and retro-Michaelreaction rates in the cracking reactor, the cracking reaction mixtureshould have an aqueous content lower than that of the HRU mixture. Watercontents typically below 5 wt. %, preferably below 1 wt. %, are achievedby operating the cracking reactor as a single stage unit, that is, bycontinuously distilling from the cracking reactor any water carried overfrom the hydrolysis reactor bleed stream and any additional watergenerated from cracking reactions.

Additional acid may be added to the recovery units as necessary toachieve practical reaction rates; preferably acid is added by way of oneor more of the feed streams. “Residual acid catalyst” is acid catalystwhich remains present as acid in the esterification reactor bleed streamand thus is carried forward to the HRU. In the HRU, acid concentrationis preferably in the range of 3.5 to 15 wt. %, and most preferably is 5to 8 wt. %. Acid concentration in the cracking reactor is typically inthe range of 7.5 to 20 wt. %, and could be higher, e.g. up to 50%. Acidconcentration preferably is from 10 to 13 wt. %, particularly for BAproduction. The amount of heavy ends in the esterification reactor bleedstream may vary but typically is in the range of from 10 to 50 wt. % ofthe combined total of the aqueous and organic-containing feed stream.

Hydrolysis reaction temperatures range from 90° to 140° C., and arepreferably from 105° to 125° C. for efficient hydrolysis rates;temperatures greater than 140° C. may lead to thermally inducedpolymerization of alkyl acrylates and of acryloxy-bearing heavy ends,resulting in undesired product loss. The residence time required for HRUhydrolysis reaction is preferably from 0.5 to 5 hours, more preferablyfrom 0.5 to 3 hours, shorter times being more economical. Lowertemperatures, and the presence of water, also favor reduced DBEformation. Cracking reactor temperatures range from 90° to 140° C.,preferably from 110° to 125° C.; cracking pressures typically range from20 mm Hg to 200 mm Hg, although higher pressures, up to 800 mm Hg may beused. The residence time for dehydration and other reactions in thecracking reactor under these conditions is preferably from 0.5 to 3hours. For the continuous production of BA, values recoveries aremaximized with two CSTR reactors in tandem, one the HRU and the otherthe cracking reactor.

In order to prevent polymerization, an effective amount of one or morepolymerization inhibitor may be added at any step in any component ofthe process. An esterification reactor process stream typically containssufficient inhibitor to prevent polymerization in the HRU and crackingreactor. If additional inhibitor addition is required, any of a largenumber of known inhibitors may be used, for example, hydroquinone, themono-methyl ether of hydroquinone, butylated hydroxy anisole,naphthaquinone, anthranil, and derivatives of these.

The second component of the invention, the distillative component,further improves known methods of distilling crude BA and provides BAsubstantially free of AA by more efficiently handling distillate andaqueous reflux. Specifically, the new distillative method provides BA inthe BA-rich stream containing less than 2,000 ppm of AA for movingforward for subsequent conventional isolation. The method also providesan AA recycle stream containing negligible BA, specifically providing anAA recycle stream (the bottom, AA-rich phase) containing less than 10ppm, preferably less than 5 ppm, of BA. In generating the crude BA forthe new distillative component of the invention, AA and BuOH areinitially fed, line 70, along with acid catalyst, to an esterificationreactor in a molar ratio of AA to BuOH in the range of 1:1.1 to 1:1.7,preferably 1:1.25 to 1:1.45, and reacted to a conversion on AA of from60 to 95%, preferably 75 to 85%, using an acid catalyst of the mineralor sulfonic acid type previously described, or a strong acid ionexchange resin; preferably sulfuric acid is used. The reactant ratio andBA conversion provide a crude BA stream which may be processed toprovide stable “aqueous mode” operation (discussed in detail below) ofthe acrylic acid separation column. Reactor contents are maintained in aboiling state during continuous distillation of the vaporized mixture ofAA, BA, BuOH and water.

Referring to FIG. 1, vaporized mixture by line 2 from the reactor 1 iscondensed, 62, and fed to phase separator 14 (in this embodiment) toprovide the first condensate. Alternatively, the vaporized mixture maybe fed directly to the column 15 for distillation as described above. Anentrainment separator, not shown, also may be mounted on the reactor, toreduce or eliminate entrainment of acid catalyst in the vaporizedmixture, thus reducing downstream corrosion potential. Phase separator14 is particularly useful when an entrainment separator is employed,assuring an organic reflux layer return to the entrainment separator,and also as a means for providing optional aqueous stream 42 to an HRU.The first condensate comprises an organic phase primarily (i.e. morethan 50%) of BA and BuOH, with some AA, and an aqueous phase primarilyof water, with some BuOH and AA. All of both phases may be fed to theacrylic acid separation column 15 by one or more line, e.g. 43, 53, or,optionally, up to 50 wt. % of the aqueous phase may be diverted via line42 to the hydrolytic recovery unit 5 (when preferably used). Additionalbutanol optionally may be fed to the column via line 54. An overheadazeotropic mixture is distilled from the acrylic acid separation column,under the pressure, temperature, and aqueous reflux conditionspreviously described, and condensed by line 16 and condenser 63 intophase separator 18, yielding a second condensate comprising BA-richorganic phase 19 and aqueous phase 20. By “BA-rich,” or “AA-rich,” ismeant that BA, or AA, is the primary (>50 wt. %) organic component of agiven phase. Concurrently, an AA-rich bottom stream, containingnegligible BA, is withdrawn from the bottom of the acrylic acidseparation column and by line 17 returned to the esterification reactor1. The amount of the recycled aqueous stream 21 is adjusted to provideat least an 8.5:1 minimum aqueous reflux ratio in the AA separationcolumn 15 in order to maintain the column in the critical “aqueous mode”operation. In the aqueous mode operation, the AA separation columnperforms a surprisingly effective separation of AA from theBA-containing feed stream (i.e. the first condensate stream or thevaporized mixture stream fed to the column), resulting in low AA lossesin the BA distillate and consequently higher BA yield, as shown infurther detail below. A small portion, 6 to 11 wt. % of the separatedaqueous phase 20, typically is fed forward as stream 22, along with theforward feeding of the BA-rich organic phase 19 in stream 23 forsubsequent conventional isolation of final product BA. Conventionally,the aqueous reflux ratio is defined as the ratio of aqueous flowreturned to the aqueous flow taken forward, here the ratio of aqueousflow in 21 to that in 22. Maintaining the specified ratio is critical tothe efficient operation of the acrylic acid separation column in theinvention.

The acrylic acid separation column may have from 20 to 50, preferably 30to 40, trays and typically is equipped with a bottom reboiler loop (notshown) and an overhead distillate line 16 through condenser 63 to phaseseparator 18. The first condensate feed typically is fed at about the10th tray in a 40 tray column, numbered from the column bottom. Ifoptionally used, added BuOH typically is fed at the 8th or 9th tray. Thecolumn operates within the limits described previously, and preferablyat a pressure of from 90 to 135 mm Hg corresponding to a preferablebottom temperature of from 80° to 85° C. The aqueous reflux ratio duringdistillation of the overhead mixture is preferably from 8.5 to 12.5 andmost preferably from 9.5 to 10.5. The flow rate of the column bottomstream in 17 is adjusted to exceed the amount of AA in the column feedby 5 to 25 wt. % to ensure that all AA remains in the column bottom.Stream 17 typically contains 5 to 20 wt. % water, the balance beingprimarily AA and AOPA. The acrylic acid separation column, run asdescribed, provides BA substantially free of AA (<2,000 ppm) and an AAbottom stream containing negligible (<10 ppm) BA.

One of the unexpected findings in the modeling and subsequentdemonstration of the acrylic acid separation column use was that twosteady states existed at the same operating conditions (that is, at thesame feed rate, feed composition, aqueous reflux flow rate, and bottomflow rate). One steady state, referred to above as the “aqueous mode,”is critical to obtaining the very low levels of AA in the BA-rich phaseand of BA in the AA-rich bottom stream as previously described. In theaqueous mode the acrylic acid separation column runs relatively “cool,”there are substantial amounts of water in the liquid on all trays, wateris present in the bottom stream, and there is negligible BA in thebottom stream. In surprising contrast, however, there exists at the sameconditions (that is, at the same feed rate, feed composition, aqueousreflux flow rate, and bottom flow rate) a second mode, the “organicmode,” which is undesirable. In the organic mode, the acrylic acidseparation column runs about 30-35° C. hotter than in the aqueous mode,considerable amounts (>10 wt %) of BA are found in the bottom stream,and the concentration of AA in the overhead mixture of BA is at least anorder of magnitude larger than the maximum of 2,000 ppm AA achieved viaaqueous mode operation. Also in the undesirable organic mode, the columnnot only is hotter than in the aqueous mode, but all the water isconcentrated in the top several trays and the bottom stream issubstantially dry. Examples 1-6 and the below-described modeling studiesprovide further detail of the unexpected finding of these modes and therationale behind running the acrylic acid separation column as defined.

Finally, there is provided a most preferred continuous process,employing all components of the invention in combination, for producingBA substantially free of acrylic acid (AA), and for recovering AA, BA,n-butanol (BuOH) and water from an esterification reactor mixturecontaining AA, BA, BuOH, water, heavy ends, and acid catalyst, whichcomprises the following steps:

a) feeding to an esterification reactor AA and BuOH, in a molar ratio offrom 1 to 1.1 to 1 to 1.7, and the acid catalyst;

b) Reacting the AA and BuOH to yield BA in a conversion of at least 60%on AA, and yielding the esterification reaction mixture comprising AA,BA, BuOH, water, heavy ends, and acid catalyst;

c) withdrawing a reactor bleed stream from the continuously convertingesterification reactor mixture while concurrently distilling AA, BA,BuOH and water from the esterification reaction mixture;

d) feeding a total aqueous and organic feed stream comprising thereactor bleed stream, water, optionally a strong acid selected from amineral acid or sulfonic acid, and optionally additional heavy ends, toa hydrolysis reactor maintained at 90° to 140° C., 50 to 1000 mm Hgpressure, and a residence time of 0.5 to 20 hours based on the totalaqueous and organic feed stream;

e) distilling an overhead stream containing AA, BA, BuOH, and water fromthe hydrolysis reactor while maintaining a hydrolysis reactor liquidconcentration of from 5 to 40 weight % water and at least 1 weight %acid, the acid comprising the acid catalyst and the optional strongacid;

f) condensing the overhead stream;

g) separating from the condensed overhead stream an organic phasecomprising BA, BuOH, and AA, and an aqueous phase comprising primarilywater, and AA, and BuOH;

h) feeding the separated organic phase to the esterification reactor;

i) feeding the separated aqueous phase to the hydrolysis reactor;

j) withdrawing from the hydrolysis reactor from 20 to 70 weight %, basedon the total aqueous and organic feed stream, of a hydrolysis reactorbleed stream;

k) feeding up to 100% of the hydrolysis reactor bleed stream to acracking reactor maintained at 90° to 140° C., a pressure of from 20 to200 mm Hg, and a residence time of 0.5 to 20 hours based on the fedreactor bleed stream;

l) distilling from the cracking reactor a cracking reactor overheadstream comprising AA, BA, BuOH, and water while maintaining a crackingreactor liquid concentration of at least 7.5 weight % acid;

m) condensing the cracking reactor overhead stream;

n) recycling to the esterification reactor the condensed crackingreactor overhead stream comprising AA, BA, BuOH, and water;

o) distilling from the esterificafion reactor, concurrently with abovesteps c) through n), a vaporized mixture comprising AA, BA, BuOH, andwater;

p) condensing the vaporized mixture to provide a first condensatecomprising an organic phase and an aqueous phase;

q) returning from 0 to 30 percent of the organic phase to an entrainmentseparator surmounting the esterification reactor; and

r) feeding from 70 to 100 percent of the organic phase and from 50 to100 percent of the aqueous phase to an acrylic acid separation column;

s) distilling from the acrylic acid separation column, at a pressure offrom 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratioof 8.5:1 to 17:1, an overhead mixture comprising an azeotropic mixtureof butanol, butyl acrylate and water;

t) removing from the distillation column an acrylic acid-rich bottomstream;

u) recycling the acrylic acid-rich bottom stream from the acrylic acidseparation column to the esterification reactor;

v) condensing the overhead mixture to provide a second condensate;

w) separating the second condensate into a butyl acrylate-rich organicphase and a separated aqueous phase; and

x) removing the butyl acrylate-rich organic phase substantially free ofAA.

This method described immediately above also may be carried out whereinsteps p), q), and r) are bypassed and 100 percent of the vaporizedmixture is fed directly to the acrylic acid separation column of step s)and distilled thereafter as described. When the vaporized mixture is feddirectly to the column, the aqueous reflux ratio is tightened to 13:1 to17:1; all other steps are identical, except there is, of course, no“first condensate.”

In the continuous methods described immediately above, the acid catalystmay be selected from sulfuric acid, a sulfonic acid, preferablymethane-, benzene-, and toluene-sulfonic acid, or a strong acid ionexchange resin. Sulfuric acid is preferred for use both as the acidcatalyst and as the optionally added mineral acid. A preferable pressurerange for carrying out the distillation in the AA separation column isfrom 90 mm to 135 mm Hg. A preferred aqueous reflux ratio is, again,from 8.5 to 12.5. The total aqueous and organic feed stream may be fedeither to a hydrolysis reactor which is a multi-plate reactivedistillation column or, preferably, to a CSTR, as previously described,thus providing hydrolytic reaction under continuously mixed conditions.The additional heavy ends here also may comprise up to 80 wt. % of thetotal aqueous and organic feed stream.

Returning to FIG. 1, streams of 20 and 19, are taken forward in 23 orseparately, as 22 and 23, and product BA is then isolated byconventional means. Thus, the process from this point forward may becompleted conventionally, for example, by feeding streams 22 and 23, toa separator where the stream is caustic-neutralized and any resulting AAsalt extracted by water. The AA-free organic phase is then dehydratedthrough a distillation column, removing final traces of water. In a nextcolumn unreacted BuOH is recovered from the overhead as its azeotropewith BA for recycle to the esterification reactor (stream 52), andpassing the bottom stream, containing substantially pure BA andinhibitors, to a product final distillation column. In this finalcolumn, pure BA is distilled overhead in a conventional manner and ableed stream containing process inhibitors is removed from the bottomfor reuse. Representative purity of the BA obtained from the processjust described typically exceeds 99.8% BA.

EXAMPLES

General Materials: AA, crude and pure n-butyl acrylate (BA), n-butanol(BuOH), and heavy end streams were obtained from plant productionstreams where indicated and were of the quality/purity indicated.Commercial polymerization inhibitors were used as purchased at levelsindicated and included hydroquinone (HQ), HQ methyl ether (MEHQ), andphenothiazine (PTZ). Heavy end components in the Examples include thefollowing materials: AOPA, butyl β-butoxy-propionate (“BBBP”), butylβ-hydroxy-propionate (“BBHP”), butoxy AOPA (“BAOPA”), butyl maleate andDBE.

Abbreviations: These include, in addition to those already defined, thefollowing terms: additional (add'l); aqueous (aq.); Comparative (Comp.);Example (Ex.); Figure (Fig.); gram (g); grams per hour (g/hr); kilogram(Kg); hour(s) (hr(s)); heavy ends or heavy end streams (“heavies”);weight (wt.); millimeters of mercury pressure (mm Hg); millimoles(mmoles or mm); pounds (lbs); vaporized mixture (vap. mixt.); roundbottom flask (r.b. flask); less than (<); more than (>); point (pt.);steady state (s.s.). In FIG. 2, data points are abbreviated as follows:open box points are in the aqueous s.s. (mode), triangles are in theorganic s.s., and circled data points are experimental/example runs asnumbered.

Analyses: Standard methods were used for determination of water;monomer, BuOH, and residual impurity and heavy end levels weredetermined by gas/liquid chromatography (GLC) on a Varian Model 3700chromatograph, using flame ionization detection. Sulfuric aciddeterminations were obtained using an Orion Research Ion analyzer pHprobe and alcoholic tetrabutylammonium hydroxide titrant. Unlessotherwise noted, H₂SO₄ concentrations given in examples are thesetitrated values. Percentages are in wt. %, unless otherwise indicated.

Values Recoveries: Recovered “values” were calculated and measured asfollows from representative heavy ends produced, for example, in BAproduction. Since all heavy ends related to BA production ultimately arederived from AA and BuOH, the values recovery data were calculated toreflect the recovery of these reactants, even though some recovery is inthe form of product BA. For example, 100 moles of BBBP contains theequivalent of 100 moles of acrylic acid and 200 moles of BuOH.Similarly, 100 moles of BAOPA contains the equivalent of 100 moles ofBuOH and 200 moles of AA. The “heavies mixture” (which is residueunaccounted for) is assumed, for weight calculation purposes, to be a1:1 molar mixture of acrylic acid and BuOH with a molecular weight of146 g/mole. BA monomer contains equivalent molar amounts of AA and BuOH.So-called “free” values are simply the same values in “free” (notincorporated as heavy end) form. Following is a list of BuOH and AAvalues for representative, characterized heavy ends from exemplified BAesterification reaction.

Component Values butyl-β-butoxy propionate (BBBP) 2 BuOH, 1 AAbutyl-β-hydroxy propionate (BBHP) 1 BuOH, 1 AA butyl-acryloxypropionate(BAOPA) 1 BuOH, 2 AA Acryloxypropionic Acid (AOPA) 2 AA n-butyl maleate1 BuOH Butylhydrogen sulfate 1 BuOH Heavies Mixture 1 BuOH, 1 AA

Process Yield: Process yield was calculated in the following manner. TheAA and BuOH present in any additional streams fed to the HRU weretreated in yield calculations as if they were fresh (i.e., raw material)AA and BuOH as fed to the esterification reactor. The BA monomer presentin additional streams fed to the HRU was treated in yield calculationsas if it were recycled BA from a downstream separation (i.e., recycledor supplemental streams); i.e. no yield increment was credited for anyrecycled BA. The yield on AA or BA can, therefore, exceed 100% whenvalues (as described above) were recovered from the HRU- andHRU/cracker-treated heavy end streams, as described. Thus, in yieldcalculation summary:${\% \quad {Yield}\quad {of}\quad {BA}},{{{based}\quad {on}\quad {AA}} = \frac{\begin{matrix}{{{mole}\quad {BA}\quad \left( {{vap}.{mixt}.} \right)} -} \\{{{mole}\quad {BA}\quad ({recycled})} - {{mole}\quad {BA}\quad \left( {{{add}'}l\quad {streams}} \right)}}\end{matrix}}{{{mole}\quad {AA}\quad \left( {{fresh}\quad {to}\quad {reactor}} \right)} + {{mole}\quad {AA}\quad \left( {{{add}'}l\quad {streams}} \right)}}}$and${\% \quad {Yield}\quad {of}\quad {BA}},{{{based}\quad {on}\quad {BuOH}} = \frac{\begin{matrix}{{{mole}\quad {BA}\quad \left( {{vap}.{mixt}.} \right)} -} \\{{{mole}\quad {BA}\quad ({recycled})} - {{mole}\quad {BA}\quad \left( {{{add}'}l\quad {streams}} \right)}}\end{matrix}}{{{mole}\quad {BuOH}\quad \left( {{fresh}\quad {to}\quad {reactor}} \right)} + {{mole}\quad {BuOH}\quad \left( {{{add}'}l\quad {streams}} \right)}}}$

Equipment: In the following Examples, HRU 5 was a 1 liter 4-neck r.b.flask equipped with a stirrer, water cooled distillation head having atake-off port leading to a 250 ml fraction cutter, phase separator 30.The HRU further was equipped with feed inlet ports 3 and 4 for reactorbleed stream and heavy end stream and other stream additions; aHastelloy dip tube of 6 mm O.D. connected by a line, 9, to crackingreactor 10 (when used) or to a bleed reservoir by line 52. The HRU washeated with a heating mantle and was mechanically stirred. The variousfeed and reflux and bleed streams were pumped into and from the reactorfrom glass feed funnels using metering pumps. HRU thermal control wasregulated by an electronic temperature controller attached to acalibrated thermocouple. All process stream lines exposed to streamscontaining sulfuric acid were constructed of Hastelloy C™ orpoly-tetrafluouroethylene (PTFE).

The cracking reactor 10 consisted of a 500 ml flask configured similarlyto the HRU regarding temperature control and process stream lines. Bleedstream inlet 9 from the HRU fed the cracking reactor via feed port andpump. Receiver 31 was a 125 ml fraction cutter.

The acrylic acid separation columns are described in specific Examples.

All percentages are by weight, based on the weight of the mixture inwhich a stated component is contained, unless otherwise indicated.

Modeling Experiments for the Acrylic Acid Separation Column:

Modeling studies were performed using “Aspen Plus,”™ an advancedflowsheet simulator from Aspen Technology, Inc. All data points wereobtained using an “Aspen” column model which had 13 theoretical traysplus reboiler and decanter and operated at an overhead pressure of 75 mmHg. The feed tray was the 4th theoretical tray from the bottom; thecolumn bottom stream was sized to contain 90 wt % AA and 10% waterduring aqueous operation. FIG. 2 shows the two “steady states” (that is,the desired “aqueous” and the undesirable “organic” modes, describedpreviously) of the acrylic acid separation column for the feeds shown inTable 1, corresponding to a reactor conversion of AA to BA of 80% at amolar ratio of AA to BuOH of 1 to 1.35. The data of FIG. 2 are plottedin concentration of AA in the organic distillate as a function of theaqueous reflux flow rate in the column.

The simulations indicated that the minimum aqueous reflux flow rateneeded to operate the acrylic acid separation column at the desiredaqueous steady state with the feed of Table 1A was approximately 15546kg (32000 lbs)/hr. The location of the aqueous/organic transition wasestimated by recognizing that the separation of BA and AA in the AAseparation column is achieved through azeotropic distillation of BAusing water as an azeotroping agent. Tables 1A and 1B below illustratehow the minimum amount of water necessary to azeotrope all the BA in theacrylic acid separation column feed was calculated. The first azeotropeto act in the column is the lowest-boiling BA/butanol/water ternaryazeotrope which at a pressure of 100 mm Hg boils at 46.4° C. andcontains 36.0% BA, 26.4% BuOH, and 37.6% water. This azeotrope depletesthe butanol in the feed and takes 10429 kg (22944 lbs)/hr of BA overheadout of a total of 20315 kg (44692 lbs)/hr present in the feed. Theamount of water needed to satisfy this first azeotrope exceeds theamount in the feed by 8198 kg (18036 lbs)/hr. Once butanol has beendepleted, the next lowest-boiling azeotrope acting in the column is theBA/water binary azeotrope which at a pressure of 100 mm Hg boils at47.6° C. and contains 61.0% BA and 39.0% water. This second azeotropetakes the remaining 9885 kg (21748 lbs)/hr of BA overhead using 6320 kg(13904 lbs)/hr of water to satisfy the azeotrope composition. Thecombined analysis of the two azeotropes shows that the total amount ofwater needed to azeotrope all the BA in the feed exceeds the amountpresent in the aqueous feed by 6320 kg (31940 lbs)/hr. This correspondsto the minimum amount of water that must be supplied via the aqueousreflux to take all the BA overhead and thus achieve aqueous modeoperation. Excellent agreement is shown between this estimate and thelocation of the aqueous/organic transition predicted by the data in FIG.2.

TABLES 1A and B Modeling Feed Conditions and Calculations 1A: AcrylicAcid Separation Column Feed for Modeling Conditions, Calculated at 80%Conversion/AA:BuOH.Ratio 1:1.35 Modeling Column Feed Component Kg(lbs)/hr wt % AA  2693  (5925)  8.1 BuOH  7648 (16826)  22.9 BA 20315(44692)  60.9 H2O  2695  (5928)  8.1 Total 33351  73371 100.0 1 B:Calculation of the Minimum Water Required to Azeotrope All BA in theFeed of Table 1A Under Modeling Conditions 1st Azeotrope 2nd AzeotropeFeed (BA/BuOH/H2O) Residual 1¹ (BA/H2O) Residual 2² per hour per hourper hour per hour per hour Comp. kg (lbs) kg (lbs) wt % kg (lbs) kg(lbs) wt % kg (lbs) BA 20315 (44692) 10429 (22944) 36.0 9885 (21748)9885 (21748) 61.0 0 0 BuOH  7648 (16826)  7648 (16826) 26.4 0 0 0 0  0.00 0 H2O  2695  (5928) 10893 (23964) 37.6 −8193 (−18036) 6320 (13940)39.0 −14518 (−31940) (Note 3) (Note 3) Notes: ¹Residual 1 is the Feed ofTable 1A reduced by the 1st Azeotrope components. ²Residual 2 isResidual 1 reduced by the 2nd Azeotrope components. ³Water required toachieve the azeotropic composition of the respective azeotropicdistillations, showing deficits by the amount indicated.

1B: Calculation of the Minimum Water Required to Azeotrope All BA in theFeed of Table 1A Under Modeling Conditions 1st Azeotrope 2nd AzeotropeFeed (BA/BuOH/H2O) Residual 1¹ (BA/H2O) Residual 2² per hour per hour erhour per hour per hour Comp. kg (lbs) kg (lbs) wt % kg (lbs) kg (lbs) wt% kg (lbs) BA 20315 (44692) 10429 (22944) 36.0   9885   (21748) 9885(21748) 61.0     0     0 BuOH  7648 (16826)  7648 (16826) 26.4     0    0   0   0  0.0     0     0 H2O  2695  (5928) 10893 (23964) 37.6−8193 (−18036) 6320 (13904) 39.0 −14518 (−31940) (Note 3) (Note 3)Notes: ¹Residual 1 is the Feed of Table 1A reduced by the 1st Azeotropecomponents. ²Residual 2 is Residual 1 reduced by the 2nd Azeotropecomponents. ³Water required to achieve the azeotropic composition of therespective azeotropic distillations, showing deficits by the amountindicated.

The modeling results presented here correspond to a particular feed tothe acrylic acid separation column. However, the same analysis can beapplied to any column feed, corresponding to any particular set ofreactor conditions, to estimate the minimum aqueous reflux requirementin the column. The ability to accurately predict the minimum waterrequirement for the acrylic acid separation column based on feedcomposition alone allows the selection of an operating reflux ratio thatminimized the heat duty and diameter of the column while ensuring stableoperation in the desired aqueous mode.

In the modeling, it was possible to control which steady state, aqueousor organic, the column operated at by starting a run at an extreme point(i.e., very high reflux rates for aqueous steady state or very lowreflux rates for organic steady state) where only the targeted steadystate exists, and then moving along the branch, either by decreasing orincreasing, respectively, the reflux flow rate until the targeted pointof operation was reached. This was achieved through examination of thesensitivity of the process to key variables; in this study, the aqueousreflux flow rate was examined. The two steady state branches of FIG. 2were obtained by performing two sensitivity studies in the program. Inthe first study, the aqueous reflux flow rate was started at a very highend of 61364 kg (135000 lbs)/hr (a reflux ratio of about 40, using 1591kg (3,500 lbs)/hr as the aqueous feed forward rate) and graduallydecreased to a very low 4545 kg (10000 lbs)/hr. (a reflux ratio of about3). This study generated the lower, “aqueous branch” of FIG. 2 whichrepresented the desired aqueous mode where the levels of AA in theorganic distillate are very low. (In this program, the lowest level ofAA in the distillate (27 ppm) was achieved with the minimum amount ofreflux (ca. 15454 kg (32000 lbs)/hr, a reflux ratio of about 9,indicated to run the column in the desired steady state mode.) When thereflux flow rate became too low, the column became inoperable in anaqueous mode and, at ca. 14090 kg (31000 lbs)/hr of reflux, a sudden andvery large increase in the distillate AA level occurred. Below 14090 kg(31000 lbs)/hr, the column operated only in the organic mode and the twomodes converged to a single solution.

In a second sensitivity study, the aqueous reflux flow rate was startedat the low end, at 4545 kg (10000 lbs)/hr, and gradually increased toca. 61364 kg (135000 lbs)/hr. This study generated the upper, “organicbranch” in FIG. 2 and represented the undesired organic mode where thelevels of AA in the organic distillate are much higher, as indicated.Moving successively along this branch (points 4-7) to above ca. 54545 kg(120000 lbs)/hr where there is enough water to force the column intoaqueous mode of operation, both branches converged to a single aqueoussteady state. Within the aqueous branch, the operating region in thissimulated study leading to BA having substantially no AA (a target of2,000 ppm, preferably <1000 ppm,AA), is a small region in the bottomaqueous branch of FIG. 2. The program also predicted high levels of BA(e.g. 23-74 wt. %) in the bottom stream when the column ran in theorganic mode. Recycle of BA to the esterification reactor is undesirablebecause it depresses the rate of conversion of AA and BuOH.

In subsequent modeling of the two steady states in the AA separatingcolumn, it was determined that bypassing the reactor condenser and phaseseparator 14 and feeding a vaporized mixture directly to the column hadthe advantage of reducing the steam duty requirement of the column.However, because the water in the feed is already vaporized, it isessentially unavailable to form an azeotrope with BA and more refluxwater is required to make up for this deficiency. For a vapor feed tothe column, the aqueous/organic transition point in FIG. 2 moves towardthe right by the amount of water in the feed, and the aqueous refluxratio range for aqueous-mode operation is tightened to 13:1 to 17:1.

Modeling also showed that refluxing any portion of the organic phase isdetrimental to column operation because any BA and butanol returned tothe column via an organic reflux will simply need to be removed again byazeotropic distillation with additional water. In addition, AA in theorganic reflux is returned to the column at the very top, leaving notrays to rectify this AA contribution out of the overhead vapor. Thesefactors increase the minimum amount of water necessary to operate in theaqueous mode, reduce the width of the aqueous operating window, andraise the minimum levels of AA that can be achieved in the distillate.

Vapor-liquid equilibrium (VLE) data indicate that butanol has the effectof depressing the volatility of AA. In accordance with the VLE data,modeling shows that column feed streams that are rich in butanol givedistillate streams that are low in AA. Therefore, low conversions andhigh butanol-to-AA ratios in the reactor which yield butanol-richeffluents are favorable for the BA/AA separation and yield the wide(8.5:1-17:1) aqueous reflux operating windows, as described. In theevent that the reactor cannot be operated under the above conditions, aprovision for a separate fresh butanol stream fed directly to the AAseparation column can be made to ensure a wide aqueous mode operatingwindow independent of the reactor conditions. Fresh butanol is best fedat or slightly below the main feed. Butanol should never be fed abovethe main, AA-containing feed. (As a light component, feeding butanolabove the main feed allows it to flash overhead quickly, leaving thetrays between the mainfeed and butanol feed with little butanol tosuppress AA volatility.)

Laboratory Confirmation of the Aqueous and Organic Modes

The existence of two steady states in the AA separation column wasconfirmed experimentally in a multi-day continuous laboratory run, fromwhich Examples 1-6 and Comparative Examples 1-2 were taken. Materialflow rates in Table 1 and in the simulations that generated FIG. 2 weremodeled on a plant-scale; in the acrylic acid separation column Examplesbelow, flow rates were scaled down such that 250 Kg (550 lbs)/hr on theabove plant scale model were equivalent to 1 g/hr in this laboratoryrun. The extended run, which approximately followed the points circledin FIG. 2, started out by demonstrating continuous operation of thecolumn in the aqueous mode for various reflux flow rates, Points 1 and2. This portion of the run was followed by an intentional decrease ofreflux water to drive the column to organic mode operation, Points 3 and4. Subsequent changes of boil-up conditions then were imposed to restorethe column to aqueous operation. In FIG. 2, points 1 and 5, 2 and 4, 6and 8, represent pairs of matching points, i.e., points of equal refluxflow rate in the aqueous and organic modes, respectively. The aqueousmode, once achieved, was maintained by a reflux ratio of from 8.5:1 to17:1, and yielded distilled BA having the desired level of AA, <2,000ppm and also a separated aqueous AA stream having substantially no BA.The measured levels of AA in BA at points 1 and 2 were 950 ppm and 200ppm AA, respectively, and of BA in AA, none (<1 ppm) was measured.

Example 1

Aqueous Mode Operation at Aqueous Reflux Ratio of 16 (Point 1 of FIG. 2)

The acrylic acid separation column 15 was a 30-tray, 1-inch diameter,Oldershaw fractional distillation column equipped with a glass condenserin line 16 and a stainless steel steam reboiler. The column was operatedat an overhead pressure of 75 mm Hg. The acrylic acid separation columnwas fed per hour with 10.8 g (0.15 mol) of AA, 30.6 g (0.41 mol) ofbutanol, 82.4 g (0.64 mol) of BA, and 12.0 g (0.67 mol) of water. Thismixture composition corresponded to a reactor condensate generated in asystem where the reactor 1 operated at an AA-to-BuOH ratio of 1:1.35 andat a conversion of 80% on AA while receiving per hour 0.18 g of recycledBA per gram of unreacted BuOH and 0.11 g of recycled water per gram ofunreacted AA. The feed tray was the 10th tray from the bottom. Theoverhead mixture distilled at a temperature of 43.5° C. and wascondensed and separated into two phases in the receiver 18. Of theBA-rich organic phase 19, 117.3 g/h (grams per hour) were collectedwhich contained, by weight, 70.3% BA, 26.0% butanol, 3.6% water, and0.1% AA. Of the separated aqueous phase 20, 110.2 g/h (94.3% of thephase) was recycled to the top of the column through line 21 and 6.7 g/h(5.7%) was moved forward through line 22, yielding an aqueous refluxratio of 16.4. The aqueous phase contained 96.6% water, 3.2% butanol,0.2% BA, and 354 ppm AA. Of the AA-rich bottom product, stream 17, 12.0g/h were collected containing 89.2% AA and 10.8% water. The resultingbottom temperature was 60.0° C. Ex. 1 corresponded to point 1 in FIG. 2.

Example 2

Aqueous Mode Operation at Aqueous Reflux Ratio of 11 (Pt. 2 of FIG. 2)

The apparatus, feed rate, feed composition, feed location, columnpressure and general column operation were the same as those describedin Ex. 1. Steam to the reboiler and aqueous condensate return rate werereduced in order to reduce the aqueous reflux flow to the column. Theoverhead product in line 16 obtained at a temperature of 42.6° C. wascondensed and separated into two phases in receiver 18. 117.1 g/h of theBA-rich organic phase 19 were collected which contained 70.4% BA, 26.0%butanol, 3.6% water, and 218 ppm of AA. Of the separated aqueous phase20, 77.2 g/h (91.8%) were recycled to the top of the column through line21 and 6.9 g/h (8.2%) were moved forward through line 22, yielding anaqueous reflux ratio of 11.2. The aqueous phase contained 96.6% water,3.2% butanol, 0.2% BA, and 81 ppm AA. 12.0 g/h of the bottom product byline 17 were collected containing 90.0% AA and 10.0% water. Theresulting bottom temperature was 60.420 C. This Example corresponded topoint 2 in FIG. 2.

Comparative Example 1

Organic-mode operation (Pt. 3 of FIG. 2)

The apparatus, feed rate, feed composition, feed location, and columnpressure and general column operation were the same as those describedin Ex. 1 and the column was initially operated in a fashion identical toEx. 2. Steam to the reboiler and aqueous condensate return rate werethen reduced in order to further reduce the aqueous reflux flow to thecolumn. The overhead product through 16 obtained at a temperature of50.1° C. was condensed and separated into two phases in receiver 18.118.5 g/h of the BA-rich organic phase 19 were collected which contained62.6% BA, 25.6% butanol, 6.1% AA, and 5.7% water. Of the separatedaqueous phase 20, 36.0 g/h (86.2%) were recycled to the top of thecolumn through line 21 and 5.8 g/h (13.8%) were moved forward throughline 22, yielding an aqueous reflux ratio of 6.3. The aqueous phasecontained 94.4% water, 3.1% butanol, 2.2% AA, and 0.3% BA. 11.8 g/h ofthe bottom product via 17 were collected containing 70.7% BA, 29.1% AAand 0.2% butanol. The resulting bottom temperature was 88.3° C. ThisExample corresponded to point 3 in FIG. 2 and demonstrated thatoperating the column at a reflux ratio below the reflux ratio range ofthe present invention leads to undesired organic-mode operation. Theresults included high levels of AA in the BA-rich organic phase 19, highlevels of BA in the bottom stream 17 and high column temperaturesrelative to the aqueous mode conditions of Examples 1-2.

Comparative Example 2

Confirmation of Two Steady States

The apparatus, feed rate, feed composition, feed location, and columnpressure were the same as those described in Ex. 1 and the column wasinitially operated under conditions identical to the completion of Comp.Ex. 1. Steam to the reboiler and aqueous condensate return rate werethen increased in order to raise the aqueous reflux to the column to thesame flow rate as in Ex. 2 (point 2 in FIG. 1). The overhead productthrough line 16 obtained at a temperature of 43.9° C. was condensed andseparated into two phases in the receiver 18. 117.9 g/h of the BA-richorganic phase 19 were collected which contained 63.9% BA, 25.8% butanol,5.2% water, and 5.1% AA. Of the aqueous phase 20, 77.2 g/h (92.5%) wererecycled to the top of the column through line 21 and 6.2 g/h (7.5%)were moved forward through line 22, yielding an aqueous reflux ratio of12.4. The aqueous phase contained 94.7% water, 3.1% butanol, 1.9% AA,and 0.3% BA. 11.9 g/h of the bottom product through 17 were collectedcontaining 60.4% BA, 39.2% AA and 0.4% butanol. The resulting bottomtemperature was 88.3° C. This Comparative. Ex. corresponded to point 4in FIG. 2 and showed that even with a reflux flow rate of 77.2 g/hr, thesame as in Example 2, and an aqueous reflux ratio of 12.4, the columnremained in the undesired state of organic mode operation and gave highlevels of AA in the organic distillate and of BA in the bottom stream,and high column temperatures, relative to results under aqueous modeconditions of Examples 1 and 2.

By demonstrating the existence of Point 4 in FIG. 2, the organic modepoint analogous to aqueous mode point 2, this Comparative Exampledemonstrated that two steady states indeed exist in the column aspredicted by the modeling described above. This Comparative Example alsodemonstrated that the two steady state branches form a “hysteresis loop”and that once the column is operating in the undesired organic mode,with sufficient heat input it remained in that mode of operation evenafter the aqueous reflux ratio rate has been increased to a leveleffective in aqueous mode operation.

Example 3

Restoration of Aqueous Mode Operation from Organic Mode Operation

The apparatus, feed rate, feed composition, feed location, and columnpressure were the same as described in Ex. 1. The column was initiallyoperated at point 3 of FIG. 2 in a run identical to Comp. Ex. 1. To thetop tray of the column was then added a stream of water at a rate of41.2 g/h. Combined with the original 36.0 g/h of aqueous reflux, thisadditional water stream provided an effective reflux flow to the columnof 77.2 g/h, the same reflux rate as in Example 2 and Comp. Ex. 2, i.e.,points 2 and 4, respectively, in FIG. 2. Reboiler steam input wasmaintained at the same level as in Comp. Ex. 1 (point 3 of FIG. 2).Through thermocouples placed in alternate trays, a cool front, primarilyof liquid water, was observed to move down the column, starting at thetop tray and descending one tray at a time until it eventually reachedthe reboiler. Thus, with no additional steam provided to the reboiler tohandle the higher load, the additional water fed to the top tray behavedas expected, in providing a cooling effect to all trays. Once the coolfront reached the reboiler, indicated by a sharp temperature drop from88.3° C. to 57.0° C., the additional fresh water stream to the top traywas discontinued and reboiler steam flow rates were increased to raisethe aqueous reflux rate from 36.0 g/h to 77.2 g/h, and the column wasallowed to reach steady state, now in the aqueous mode, at the higherreflux rate.

The overhead mixture through line 16 obtained at a temperature of 42.6°C. was condensed and separated into two phases in receiver 18; 117.0 g/hof the BA-rich organic phase 19 were collected which contained 70.5% BA,26.0% butanol, 3.5% water, and 263 ppm of AA. Of the separated aqueousphase 20, 77.2 g/h (91.8%) were recycled to the top of the columnthrough line 21 and 6.9 g/h (8.2%) were moved forward through line 22,yielding an aqueous reflux ratio of 11.2. The aqueous phase contained96.6% water, 3.2% butanol, 0.2% BA, and 75 ppm AA. 12.1 g/h of the AArich bottom stream 17 were collected containing 89.5 weight % AA and10.5 weight % water. The resulting bottom temperature was 60.2° C. Thisoutcome corresponded to point 2 in FIG. 2 and was substantiallyidentical to that of Ex. 2. Thus, Ex. 3 demonstrated a short-cut methodto return the column to the desired aqueous mode operation from a pointon the organic mode branch. In the aqueous mode, the acrylic acidseparation column is run with water in all trays and in the bottomstream 17 while in the undesired organic mode, water concentrates in thetop several trays and bottom stream 17 is devoid of water. Although inthis Example the acrylic acid separation column started in the organicmode, by the treatment shown the column was made operable in the desiredaqueous mode. This result was especially important in view of thefindings of Comp. Ex. 2 which had confirmed that the two steady statesin this particular system for producing BA form the “hysteresis loop” asshown in FIG. 2.

Example 4

Aqueous Mode Operation at Aqueous Reflux Ratio of 9.6

Using apparatus described in Example 1, the column was operated at anoverhead pressure of 75 mm Hg and was fed per hour with 5.6 g (0.08 mol)of AA, 34.8 g (0.47 mol) of butanol, 96.4 g (0.75 mol) of BA, and 13.1 g(0.73 mol) of water. This mixture composition corresponded to a reactorvaporized mixture generated in BA esterification wherein the reactor 1operated at an AA-to-butanol ratio of 1:1.5 and conversion of 90% on AA,recycling 18% of BA per unit wt. of unreacted butanol via stream 52 and7% of water per unit wt. of unreacted AA via stream 17. The feed traywas the 10th tray from the bottom. The overhead distillate 16 obtainedat a temperature of 42.2° C. was condensed and separated into two phasesin the receiver 18. 135.8 g/h of the BA-rich organic phase 19 werecollected; it contained 71.0 weight % BA, 25.5 weight % butanol, 3.5weight % water, and 550 ppm of AA. Of the aqueous phase 20, 78.9 g/h(90.6%) were recycled to the top of the column through stream 21 and 8.2g/h (9.4%) were moved forward through stream 22, thus providing anaqueous reflux ratio of 9.6. The separated aqueous phase contained 96.7weight % water, 3.1 weight % butanol, 0.2 weight % BA, and 209 ppm AA.6.0 g/h of the AA-rich bottom stream by line 17 were collectedcontaining 93.1 weight % AA and 6.9 weight % water.

Example 5

Aqueous Mode Operation at Aqueous Reflux Ratio of 11.0

The apparatus, feed rate, feed composition, feed location, and columnpressure were the same as in Example 4. The overhead mixture in 16obtained at a temperature of 41.9° C. was condensed and separated intotwo phases in receiver 18. 135.9 g/h of the BA-rich organic phase 19were collected and contained 70.9 weight % BA, 25.4 weight % butanol,3.5 weight % water, and 779 ppm of AA. Of the aqueous phase 20, 89.5 g/h(91.6%) were recycled back to the top of the column through stream 21and 8.2 g/h (8.4%) were moved forward through stream 22, for an aqueousreflux ratio of 11.0. The aqueous phase contained 96.7 weight % water,3.1 weight % butanol, 0.2 weight % BA, and 286 ppm AA. 6.0 g/h of theAA-rich bottom stream 17 were collected containing 92.8 weight % AA and7.2 weight % water. This Example demonstrated an increase of AA in theBA-rich organic phase from 550 ppm to 779 ppm, under these conditions,as the amount of aqueous reflux increased relative to that of Example 4(9.6 reflux ratio).

Example 6

Aqueous Mode Operation at Aqueous Reflux Ratio of 9.7 with a 35-trayColumn

A five-tray section was added to the apparatus used in Example 1, thusproviding a 35-tray, 1-inch diameter, Oldershaw fractional distillationcolumn equipped with a glass condenser and stainless steel steamreboiler. The feed rate, feed composition, feed location, and columnpressure were the same as those of Example 4. The overhead mixture in 16obtained at a temperature of 42.2° C. was condensed and separated intotwo phases in receiver 18. 135.8 g/h of the BA-rich organic phase 19were collected which contained 71.0 weight % BA, 25.5 weight % butanol,3.5 weight % water, and 193 ppm of AA. Of the separated aqueous phase20, 78.9 g/h (90.7%) were recycled to the top of the column throughstream 21 and 8.1 g/h (9.3%) were moved forward through stream 22, foran aqueous reflux ratio of 9.7. The aqueous phase contained 96.7 weight% water, 3.1 weight % butanol, 0.2 weight % BA, and 72 ppm AA. 6.1 g/hof the AA-rich bottom stream 17 were collected containing 92.5 weight %AA and 7.5 weight % water. The resulting bottom temperature was 62.3° C.This Example demonstrated that adding 5 trays to the rectifying sectionof the AA separation column further reduced AA in the BA-rich organicphase from 550 ppm in Ex. 4 to 193 ppm.

Comparative Example 3

Cracking Reactor Processing Without Use of an HRU

This comparative example was performed in the above described 500 mlcracking reactor, using feed streams described and without use of a HRU.Thus, 73.46 g/hr of a feed containing the composition listed in table 2was fed to a CFSTR maintained at 130° C., 35 mm Hg pressure, 60 min.residence time, and a catalyst concentration of 8.07 wt % H₂SO₄. A totalof 55.24 g/hr of a single phase distillate was recovered with thecomposition listed in table 3. A bleed stream of 18.22 g/hr was bledfrom the cracking reactor and discarded as waste oil. The AA and BuOHrecovered values are summarized in tables 10 and 11, which shows that,after recovery of free values, only 15.0% of the AA values in heaviesand 11.6% of BuOH values in heavies were recovered.

TABLE 2 Feed Stream Composition for Comparative Example 3 Feed Streamg/hr in mmol/hr in Feed mmol/hr Values Components Feed Stream mm Bu mmAA BuOH 0.70 9 9 0 BA 31.66 247 247 247 AA 15.65 217 0 217 BBBP 1.47 714 7 BBHP 0.41 3 3 3 BAOPA 3.89 19 19 38 AOPA 3.60 25 0 50 Butyl Maleate1.47 9 9 0 BuOSO₂OH 2.31* 15 15 0 “Heavies Mixture” 10.10 69 69 69Inhibitor 2.20 NA Totals 73.46 385 631 *1.47 g/hr calculated as H₂SO₄

TABLE 3 Distilled Overhead Stream Composition of Comparative Example 3Feed Stream mmol/hr Values Components g/hr mmol/hr mm Bu mm AA BuOh0.884 12 12 BA 33.199 259 259 259 AA 16.572 230 0 230 BBBP 0.718 4 8 4Water 0.829 46 Dibutyl Ether 0.017 0.13 0.13 0 High Boilers 3.022 NATotals 55.24 271 489

Example 7

HRU Evaluation Under Effective Conditions

A total of 73.46 g of organic feed containing the composition listed inTable 4 was fed to the HRU maintained at a temperature of 108° C., 760mm Hg, 144 min. residence time, 16 wt % reactor water, and a catalystconcentration of 2.7 wt % H₂SO₄. In addition, 48.0 g/hr ofesterification reactor first aqueous distillate (comprising 93.0% H₂O,6.0% AA, and 1.0% BuOH) also was fed to the HRU to compensate for waterdistilled and removed with the organic distillate and reactor bleed andto simulate recycle of the aqueous distillate to the HRU. A total of39.35 g/hr of organic distillate and 38.27 g/hr of aqueous distillatewere collected and analyzed. The results of the analysis are summarizedin Table 5. The organic phase was separated for return to anesterification reactor, when used. A bleed stream of 43.24 g/hr was bledfrom the HRU and constituted the total feed to the bleed stripper CFSTR.

TABLE 4 Feed Stream Composition for HRU Feed, Example 7 Feed Stream g/hrin mmol/hr in Feed mmol/hr Values Components Feed Stream mm Bu mm AABuOH 0.85 11 11 0 BA 32.627 255 255 255 AA 15.471 215 0 215 BBBP 0.951 510 5 BBHP 0.410 3 3 3 BAOPA 7.320 37 37 74 AOPA 1.57 11 0 22 ButylMaleate 0.712 4 4 0 BuOSO₂OH* 2.309 15 15 — Heavies Mixture 9.034 62 6262 Inhibitor 2.20 NA Totals 73.46 397 636 *1.47 g calculated as H₂SO₄

TABLE 5 Compositions of Streams from HRU Evaluation of Example 7 Org.Aqueous Phase Organic mmol/hr Values Phase Component g/hr mmol/hr ValuesBu AA g/hr BuOH 3.959 54 54 0 1.102 BA 28.159 220 220 220 — AA 5.245 730 73 2.465 BBBP 0.014 0.07 .14 .07 — Water 1.624 90 — — 34.703 DBE 0.0280.22 .44 .44 — High Boilers 0.321 — — — — Totals 39.350 347 274 29338.270

Example 8

Cracking Reactor Evaluation Under Effective Conditions

The total of 43.24 g/hr of the HRU bleed stream from Example 7 was fedto the cracking reactor CSTR described above and held at a temperatureof 130° C. at 100 mm Hg pressure, and a residence time of 120 min. Atotal of 33.63 g of distillate, the cracking reactor overhead stream,was obtained for return to an esterification reactor. A total of 9.61g/hr of cracking reactor residue stream was collected and discarded aswaste oil. The total AA and BuOH recoveries for the combination of bothunits (HRU and cracking reactor) is summarized in Tables 10 and 11. Theresults show that, after recovery of free values, 68.7% of the AA valuesin heavies and 59.5% of the BuOH values in heavies were recovered.

TABLE 6 Composition of the Cracking Reactor Overhead Stream of Example 8Cracking Reactor Overhead Stream mmol/hr Values Components g/hr mmol/hrmm Bu mm AA BuOH  0.789  11 11  0 BA  7.572  59 59 59 AA 16.716 232  0232 Water  7.750 431 High Boilers  0.803 — — — Totals 33.630 70 291

Example 9

HRU Evaluation under More Severe Conditions and with a Cracking Reactorin Tandem

This example under more severe HRU operating conditions (higher acidconcentration) than in Ex's. 7 and 8 afforded higher recovery of AA andBuOH values. Thus, 73.46 g/hr of a feed stream with the compositionlisted in Table 7 was fed to the HRU maintained at 114° C., 760 mm Hg,5.1 wt % H₂SO₄ (7.5 wt % by mass balance (MB)), 144 min. residence time,and a water concentration of 15.5 wt %. In addition, 46.48 g of aqueousfeed comprising 93% water, 6% AA, and 1% butanol was fed to the HRU tosimulate recycle of the aqueous distillate plus makeup of water lost tothe organic distillate and bottoms bleed. A total of 41.96 g/hr of HRUorganic distillate (composition in Table 8), 38.80 g/hr of HRU aqueousdistillate, and 28.22 g/hr of cracking reactor overhead stream(composition in Table 9) were recovered and analyzed. The crackingreactor overhead stream was obtained by feeding the HRU bleed stream(39.18 g/hr) to the cracking reactor maintained at the followingconditions: 130° C., 100 mm Hg, 120 min. residence time, 26.8% H₂SO₄ (byMB). The total AA and BuOH recoveries for this tandem combination arelisted in Tables 10 and 11 and show that, after recovery of free values,81.7% of the AA values and 65.2% of the BuOH values in heavies wererecovered.

TABLE 7 Feed Stream Composition for Example 9 mmol/hr Values Componentg/hr Feed mmol/hr Feed Bu AA BuOH 0.528 7 7 0 BA 32.480 254 254 254 AA15.324 213 0 213 BBBP 0.804 4 8 4 BBHP 0.263 2 2 2 BAOPA 7.173 36 36 72AOPA 1.423 10 0 20 Butyl Maleate 0.565 4 4 0 BuHSO₄ 4.618* 30.0 30 0Heavies Mixture 8.078 55 55 55 Inhibitor 2.20 NA Totals 73.46 396 620*2.94 calculated as H₂SO₄

TABLE 8 Compositions of Streams from HRU Evaluation of Example 9 Org.phase Org mmol/hr Values Component g/hr Org. wt % mmol/hr Bu AA BuOH4.126 9.832 56 56 0 BA 30.438 72.54 238 238 238 AA 6.331 15.089 88 0 88BBBP 0.047 0.112 0.23 Water 0.985 2.347 55 DBE 0.034 0.081 0.26 HighBoilers — — — Totals 41.96 294 326

TABLE 9 Composition of the Cracking Reactor Overhead Stream of Example 9mmol/hr Values Component g/hr wt % mmol/hr Bu AA BuOH 0.391  1.387 5.28 5  0 BA 6.364 22.552 50 50  50 AA 15.566 55.161 216 216 Water 5.76520.428 320 High Boilers 0.132  0.466 — DBE 0.002 Totals 28.226 55 266

Example 10

Continuous Process for Producing Butyl Acrylate

The esterification reactor 1 was a 2 L round bottom, Pyrex, flaskequipped with a two plate (5.0 cm diameter) Oldershaw distillationcolumn (serving as an acid catalyst entrainment separator), a condenser,thermocouple, feed ports attached to appropriate fluid metering pumps,and lines leading to a hydrolytic reactor unit (HRU, 5) and crackingreactor 10, described more fully below. Reactor working capacity was 750ml of reaction mixture containing 2.50 wt % of sulfuric acid catalyst.The reaction temperature was 89° C. and the pressure was 127 mm Hg.Reactor 1 was fed with 182.90 g/hr of fresh crude AA (assay: 96% AA byweight, 2435 mmol/hr), 182.48 g/hr of fresh n-butanol (BuOH, 2466mmol/hr), and 1.71 g/hr of fresh H₂SO₄ (95.5 wt % acid). The reactor wasfed with a total of 655.3 g/hr of material composed of 223.85 g/hr AA(3105 mmol/hr), 316.90 g/hr BuOH (4282 mmol/hr), an HRU condensed andseparated overhead organic stream, a cracking reactor overheadcondensate, an AA separation column bottom stream, and a BA/BuOH/H₂Omixture representing streams from recovery and recycle of the followingstreams: (a) unreacted BuOH in a downstream BuOH/BA azeotropicdistillation column; (b) a BuOH/BA recovery stream from stripping ofwaste aqueous streams before sending the stream to waste treatment; and(c) a portion of final product distillation column bottoms. (Thesestreams comprise the typical feed and supplemental (e.g. recycle)streams used in a representative plant continuous process). The total BAthus fed to the esterification reactor from these sources was 88.07 g/hr(688 mmol/hr), of which 50.85 g/hr represents recycle from downstreamseparation columns. AA and BuOH were accordingly used in a mole ratio of1:1.38.

The reactor was maintained at a residence time of approximately 60minutes whereby 749.8 g/hr of total material was distilled off as thereactor overhead distillate through the Oldershaw column, condensed, andseparated in two phases. A portion of the organic phase (160 g/hr) wasreturned to the head of the distillation column as reflux. The remaining563.8 g/hr of organic distillate containing 38.56 g/hr of AA, 124.59g/hr of BuOH, and 371.24 g/hr of BA was fed to the acrylic acidseparation column, 15. The reactor's aqueous condensed vaporized mixturewas separated (26.0 g/hr containing 2.12 g/hr of BuOH and 0.619 g/hr ofAA) and split in the following fashion: 22.4 g/hr to the AA separationcolumn via line 53 and 3.6 g/hr to the HRU, via 42.

The HRU 5 and cracking reactor 10 are identical to those described inExamples 7, 8, and 9. Accordingly, 65.5 g/hr of esterification reactorbleed stream containing 4.33 g/hr of BuOH, 6.19 g/hr of AA, 34.23 g/hrof BA, and other related high boilers and inhibitors were fed to HRU 5via line 3 and maintained at 122° C., 760 mm Hg pressure, 317 minresidence time, 6.26 wt % H₂O, and an acid catalyst concentration of7.58 wt % H₂SO₄. Additionally 3.6 g/hr of reactor aqueous condenseddistillate was fed to the HRU. From the HRU, a total of 80.5 g/hr ofmaterial was distilled as an overhead stream, condensed, and separated.The entire separated aqueous phase (38.3 g/hr, containing 2.47 g/hr ofBuOH and 1.01 g/hr of AA) was returned to the HRU as reflux via 7. Theseparated organic phase (42.1 g/hr containing 7.07 g/hr BuOH, 3.19 g/hrAA, and 29.72 g/hr BA) was returned to the esterification reactor as arecovered recycle stream via 8. A bleed stream of 27.0 g/hr was removedfrom the HRU via 9 and fed to cracking reactor 10 maintained at 120° C.,35 mm Hg pressure, 815 min residence time, and 20.5 wt % H₂SO₄. A totalof 17 g/hr (containing 0.397 g/hr BuOH, 8.44 g/hr AA, and 4.78 g/hr BA)of material was distilled off and condensed in 31. This combinedcondensate was returned via 12 to the esterification reactor as recycle.The bleed stream from the cracking reactor was discarded as waste oilvia 13.

The acrylic acid separation column 15 consisted of a 35 plate, 5.0 cmdiameter, Oldershaw distillation column equipped with a steam jacketed,stainless steel, reboiler and water cooled condenser system.Accordingly, 563.8 g/hr of esterification reactor organic layercondensed distillate and 22.4 g/hr of ester reactor aqueous layercondensed distillate (composition described above) were fed to 15operated at a head pressure of 260 mm Hg, a base temperature of 82° C.,and an aqueous reflux ratio of 9.61. A total of 907.3 g/hr of overheadmixture was obtained by distillation through the column, condensed, andseparated, in 18. A total of 400.7 g/hr of separated aqueous phase wascollected of which 363.4 g/hr was returned to the head of the column asreflux vial line 21. The BA-rich organic phase (506.60 g/hr) containingthe BA product was substantially free of AA (1450 ppm). An AA-richbottom stream of 42.3 g/hr was removed from the column (35.35 g/hr AA)and recycled via 17 to the esterification reactor. Recovery data areincluded in Tables 10 and 11.

With the entire BA process operated in this fashion, a quantitativeyield of BA on BuOH was realized and a 102.7% yield of BA on AA wasrealized (of a 104.8% theoretical yield, based on the AA and AOPAcontent of the fresh crude AA charged).

Example 11

Continuous Process for Producing BA, Including Recycle of AdditionalStreams

The esterification reactor and related process equipment utilized inthis Example is identical to that described in Example 10. Reactorworking capacity was 1000 ml of reaction mixture containing 2.25 wt % ofsulfuric acid catalyst. All working units were fed as hereafterdescribed. Reactor 1 was fed with 183.90 g/hr of fresh crude AA (assay:96% AA by weight, 2449 mmol/hr), 207.03 g/hr of fresh n-butanol (BuOH,2798 mmol/hr), and 2.05 g/hr of fresh H₂SO₄ (95.5 wt % acid). The HRUwas fed via additional streams of 30.58 g/hr (424 mmol/hr) of AA and4.11 g/hr (55.5 mmol/hr) of BuOH bringing the total fresh AA feed to thesystem to 207.12 g/hr (2873 mmol/hr) and the total fresh BuOH feed to211.14 g/hr (2853 mmol/hr). Reactor 1 thus was fed with a total of 866.9g/hr of material composed of: 260.10 g/hr AA (3607 mmol/hr), 406.6 g/hrBuOH (5495 mmol/hr), an HRU overhead condensed organic layer, a crackingreactor overhead stream, an AA separation column bottom stream, and aBA/BuOH/H₂O mixture representing streams from recovery and recycle ofthe following streams: (a) unreacted BuOH in a downstream BuOH/BAazeotropic distillation column; (b) a BuOH/BA recovery stream fromstripping of waste aqueous streams before sending the stream to wastetreatment; and (c) a portion of final product distillation columnbottoms. (These streams comprise the typical feed and recycled streamsused in a fully integrated plant continuous process). The total BA fedto the esterification reactor from these sources was 144.6 g/hr, ofwhich 84.14 g/hr represented recycled BA from downstream separationcolumns and supplemental waste streams. AA and BuOH were accordinglyused in reactor 1 in a mole ratio of 1:1.52.

The reactor was maintained at a residence time of approximately 60minutes whereby 995.9 g/hr of total material was distilled as thereactor overhead distillate through the Oldershaw distillation column,condensed, and separated in two phases. A portion of the organic phase(216.1 g/hr) was returned to the head of the distillation column asreflux. The remaining 719.8 g/hr of organic distillate containing 45.20g/hr of AA, 183.10 g/hr of BuOH, and 451.0 g/hr of BA was fed to theacrylic acid separation column, 15, which is described below. Thereactor aqueous distillate (60.10 g/hr containing 4.78 g/hr of BuOH and1.06 g/hr of AA) was split in the following fashion: 36.5 g/hr to the AAseparation column and 23.6 g/hr to the HRU.

The HRU 5 and cracking reactor 10 are identical to those described inExamples 7, 8, and 9. Accordingly 87.1 g/hr of esterification reactorbleed stream via line 3 and 128.7 g/hr of additional streams containing11.3 g/hr of BuOH, 38.3 g/hr of AA, 37.1 g/hr of BA, and other relatedhigh boilers and inhibitors via line 4 were fed to the HRU maintained at122° C., 760 mm Hg pressure, 150 min residence time, 12.8 wt % H₂O, anda catalyst concentration of 9.3 wt % H₂SO₄. Additionally 23.60 g/hr ofreactor aqueous distillate was fed to the HRU. From the HRU, a total of198.7 g/hr of material was distilled, condensed, and separated. Theentire aqueous distillate (114.6 g/hr, containing 4.35 g/hr of BuOH and6.25 g/hr of AA) was returned to the HRU as reflux via line 7. Theorganic distillate (84.1 g/hr, containing 12.2 g/hr BuOH, 9.94 g/hr AA,and 56.3 g/hr BA) was returned to the esterification reactor as recyclevia line 8. An HRU bleed stream of 155.2 g/hr was removed from the HRUbottom and fed to cracking reactor 10 maintained at 120° C., 35 mm Hgpressure, 180 min residence time, and 24.0 wt % H₂SO₄. A total of 70.6g/hr (containing 2.00 g/hr BuOH, 34.3 g/hr AA, and 9.29 g/hr BA) ofcracking reactor overhead stream was distilled and condensed. Thiscondensate was recovered and returned to the main esterification reactoras a recycled stream and the cracking reactor bottom bleed stream wasdiscarded as waste oil.

The acrylic acid separation column 15 consisted of a 35 plate, 5.0 cmdiameter, Oldershaw distillation column equipped with a steam jacketed,stainless steel, reboiler and water cooled condenser system.Accordingly, 719.8 g/hr of esterification reactor organic distillatecondensate and 36.5 g/hr of reactor aqueous distillate condensate(compositions as described above) were fed to the acrylic acidseparation column operated at a head pressure of 260 mm Hg, a basetemperature of 82° C., and an aqueous reflux ratio of 12.5. A total of1193.8 g/hr of overhead mixture was obtained by distillation through theAA separation column, condensed, and separated. A total of 526.4 g/hr ofseparated aqueous phase was collected of which 487.9 g/hr was returnedto the head of the AA separation column as reflux via line 21, thebalance of the aqueous phase moved forward with the remainder of thecondensate. The BA-rich organic phase (667.4 g/hr) containing the BAproduct was also moved forward for further isolation; it wassubstantially free of AA, containing 1500 ppm AA. An AA-rich bottomstream of 50.4 g/hr was removed from the AA separation column (43.56g/hr AA) and recycled via 17 to the esterification reactor. Recoverydata are included in Tables 10 and 11.

With the entire process for producing BA operated in the processdescribed by this Example, a BA yield based on BuOH was 100.5%; on AA, aBA yield of 99.8% was realized.

TABLE 10 Summary of AA Values Fed and Recovered AA Values Free AA AATotal AA Recovered Total AA % AA Values Values In Values From ValuesRecovery Total Example Fed Heavies Fed Heavies Recovered From % AANumber (mmol) (mmol) (mmol) (mmol) (mmol) Heavies Recovery Comp. Ex. 3464 167 631  25 489 15.0% 77.5% Ex.'s 7-8 470 166 636 114 584 68.7%91.8% Ex. 9 467 153 620 125 592 81.7% 95.5% 10 354 130 484 430  77 59.2%89.1% 11 822 764 1586  1127  305 39.9% 71.1%

TABLE 11 Summary of BuOH Values Fed and Recovered BuOH Values Free BuOHBuOH Total BuOH Recovered Total BuOH % BuOH Values Values In Values FromValues Recovery Total Example Fed Heavies Fed Heavies Recovered From %BuOH Number (mmol) (mmol) (mmol) (mmol) (mmol) Heavies Recovery Comp.Ex. 3 256 129 385  15 271 11.6% 70.4% Exs. 7-8 266 131 397  78 344 59.5%86.6% Ex. 9 261 135 396  88 349 65.2% 88.1% 10 330 124 454 370  40 32.3%81.5% 11 443 471 914 705 262 55.6% 77.1%

We claim:
 1. A method of continuously recovering a C₁-C₄ alkyl acrylatesubstantially free of acrylic acid from an esterification reactionmixture, comprising the steps of: (A) distilling from an esterificationreactor a vaporized mixture comprising acrylic acid, a C₁-C₄ alkylacrylate, an C₁-C₄ alkanol, and water; (B) separating the vaporizedmixture into an organic phase and an aqueous phase; and (C) feeding atleast a portion of the organic phase and at least a portion of theaqueous phase to an acrylic acid separation column; (D) distillingoverhead from the acrylic acid separation column at an aqueous refluxratio of 8.5:1 to 17:1 an azeotropic mixture comprising a C₁-C₄ alkylacrylate, C₁-C₄ alkanol, and water; and (E) separating from theazeotropic mixture a C₁-C₄ alkyl acrylate organic phase substantiallyfree of acrylic acid.
 2. The method of claim 1, wherein from 70 to 100percent of the organic phase and 50 to 100 percent of the aqueous phaseis fed to the acrylic acid separation column.
 3. The method of claim 1,further comprising the step of: (F) recovering an acrylic acid richbottom stream from the acrylic acid separation column and recycling thebottom stream to the esterification reactor.
 4. The method of claim 1,wherein the aqueous reflux ratio is 8.5:1 to 12.5:1.
 5. The method ofclaim 1 wherein, the vaporized mixture is directly fed to the acrylicacid separation column and distilled at a reflux ratio of 13:1 to 17:1.6. A method of continuously recovering n-butyl acrylate substantiallyfree of acrylic acid from an esterification reaction mixture, comprisingthe steps of: (A) distilling from an esterification reactor a vaporizedmixture comprising acrylic acid, n-butyl acrylate, n-butanol, and water;(B) separating the vaporized mixture into an organic phase and anaqueous phase; and (C) feeding at least a portion of the organic phaseand at least a portion of the aqueous phase to an acrylic acidseparation column; (D) distilling overhead from the acrylic acidseparation column at an aqueous reflux ratio of 8.5:1 to 17:1 anazeotropic mixture comprising n-butyl acrylate, n-butanol, and water;and (E) separating from the azeotropic mixture n-butyl acrylate organicphase substantially free of acrylic acid.
 7. The method of claim 6,wherein from 70 to 100 percent of the organic phase and 50 to 100percent of the aqueous phase is fed to the acrylic acid separationcolumn.
 8. The method of claim 6, further comprising the step of: (F)recovering an acrylic acid rich bottom stream from the acrylic acidseparation column and recycling the bottom stream to the esterificationreactor.
 9. The method of claim 6, wherein the aqueous reflux ratio is8.5:1 to 12.5:1.
 10. The method of claim 6 wherein, the vaporizedmixture is directly fed to the acrylic acid separation column anddistilled at a reflux ratio of 13:1 to 17:1.